Process for increasing the yield of an isomerization zone

ABSTRACT

A process for increasing a yield of an isomerization zone by removing at least a portion of the C 6  cyclic hydrocarbons from a stream prior to it being passed into the isomerization zone. Additionally, disproportionation reaction selectivity is also increased, producing valuable C 3  hydrocarbons and C 4  hydrocarbons. Also, a higher ring opening conversion of C 5  cyclic hydrocarbons is observed. The isomerization zone may have an average operating temperature of at least 176° C. and an outlet molar ratio of hydrogen to hydrocarbon feed in the isomerization zone is less than about 0.2.

FIELD OF THE INVENTION

This invention relates to processes for separating out various fractionsof a naphtha stream, and more particularly processes which convertiso-paraffins into normal paraffins in an isomerization zone.

BACKGROUND OF THE INVENTION

Ethylene and propylene are important chemicals for use in the productionof other useful materials, such as polyethylene and polypropylene.Polyethylene and polypropylene are two of the most common plastics foundin use today and have a wide variety of uses for, for example, amaterial for fabrication and as a material for packaging. Other uses forethylene and propylene include the production of vinyl chloride,ethylene oxide, ethylbenzene and alcohol.

The great bulk of the ethylene consumed in the production of theplastics and petrochemicals such as polyethylene is produced by thethermal cracking of higher molecular weight hydrocarbons. Steam isusually mixed with the feed stream to the cracking reactor to reduce thehydrocarbon partial pressure and enhance olefin yield and to reduce theformation and deposition of carbonaceous material in the crackingreactors. The process is therefore often referred to a steam cracking orpyrolysis.

The composition of the feed to the steam cracking reactor affects theresults. A fundamental basis of this is the propensity of somehydrocarbons to crack more easily than others. The normal ranking oftendency of the hydrocarbons to crack to ethylene is normally given as:normal paraffins; iso-paraffins; olefins; naphthenes; and, aromatics.Benzene and other aromatics are particularly resistant to steam crackingand undesirable as cracking feed stocks, with only the alkyl side chainsbeing cracked to produce the desired product.

The feed stream to a steam cracking unit can be quite diverse and can bechosen from a variety of petroleum fractions. The feed stream to thesubject process preferably has a boiling point range falling within thenaphtha boiling point range or about 36 to 205° C. It is preferred thatthe feed stream does not contain appreciable amounts, e.g. more than 5mole %, of C₁₂ hydrocarbons. A representative feed stream to the subjectprocess is a C₅-C₁₁ fraction produced by fractional distillation of ahydrotreated petroleum fraction. Hydrotreating is desired to reduce thesulfur and nitrogen content of the feed down to acceptable levels. Asecond representative feed is a similar fraction comprising C₅ throughC₉ hydrocarbons.

The feed to a steam cracking unit is also normally a mixture ofhydrocarbons varying both by type of hydrocarbon and carbon number. Thisvariety results in it being very difficult to separate less desirablefeed components, such as naphthenes and aromatics, from the feed streamby fractional distillation. The hydrocarbons that are not the normalparaffins can be removed by solvent extraction or adsorption. Thesehydrocarbons can be upgraded to improve the feedstock to the steamcracking unit.

One way to upgrade these hydrocarbons is to pass the non-normalparaffins to an isomerization zone. In the isomerization zone, thenon-normal paraffins are converted, in the presence of a catalyst, intonormal paraffins.

Based upon current designs, conversion of iC₅ hydrocarbons and iC₆hydrocarbons to normal paraffins in an isomerization zone is limited, byequilibrium conditions, to about 25% and 13% per pass, respectively.Based upon typical processing conditions, full conversion of theiso-paraffins entails large recycle streams, large fractionationcolumns, and large utility costs. The per pass conversion rates can beincreased for example, by increasing the temperature of theisomerization zone, by lowering the liquid hourly space velocities(LHSV), or both, which leads to the cracking of some of paraffins tolighter C₄− hydrocarbons. The cracking reactions can lead to theproduction of undesired low value methane.

It would be desirable to have conditions for an isomerization zone inwhich the iC₅ hydrocarbon and iC₆ hydrocarbon conversions to normalparaffins are increased while maintaining low methane formation.

SUMMARY OF THE INVENTION

It has been discovered that the conversion to normal paraffins in theisomerization zone can be increased by removing, or significantlyreducing, the amount of C₆ cyclic hydrocarbons in the stream passinginto the isomerization zone. For a feed stream with C₅ and C₆hydrocarbons where at least a portion of the C₆ cyclic hydrocarbons havebeen removed, more favorable ratios of the desired normal paraffinyields over undesired methane yields are obtained in an isomerizationzone with an outlet molar ratio of hydrogen to hydrocarbon of less thanabout 0.2 preferably with a temperature of at least 176° C.

Accordingly, in a first embodiment of the invention a process forincreasing the yield of an isomerization zone is provided whichincludes: separating a portion of C₆ cyclic hydrocarbons from a naphthastream comprising C₅+ hydrocarbons to provide a C₆ cyclic hydrocarbonslean stream; separating iC₅ hydrocarbons and iC₆ hydrocarbons from theC₆ cyclic hydrocarbons lean stream; passing at least one stream beingrich in iC₅ hydrocarbons, iC₆ hydrocarbons, or both to an isomerizationzone; and, maintaining an outlet molar ratio of hydrogen to feedhydrocarbon in the isomerization zone to be less than about 0.2.

In some embodiments, an average temperature of the isomerization zone ismaintained to be to be at least approximately 176° C., or at least atleast approximately 190° C.

A weight percentage of normal paraffins in a product stream from theisomerization zone may be at least 20%. Additionally, a weightpercentage of methane in the product stream may be less that 1.0%, ormay be less than 0.5%.

A weight ratio of normal paraffins to methane in a product stream fromthe isomerization zone in one or more embodiments may be at least 75, ormay be at least 100.

In a second embodiment of the invention a process for increasing theyield of an isomerization zone is provided which includes: removing C₆cyclic hydrocarbons from at least one stream comprising iC₅hydrocarbons, iC₆ hydrocarbons, or both, to provide a C₆ cyclichydrocarbons lean stream; and, passing the C₆ cyclic hydrocarbons leanstream an isomerization zone. The isomerization zone may have anoperating temperature of at least 176° C. and an outlet molar ratio ofhydrogen to hydrocarbon in the isomerization zone less than about 0.7.

It is contemplated that the operating temperature is at least 190° C.

It is also contemplated that the outlet molar ratio of hydrogen tohydrocarbon in the isomerization zone is less than about 0.2.

In some embodiments of the present invention, a weight percentage ofnormal paraffins in the product stream from the isomerization zone is atleast 20%. Additionally, a weight percentage of methane in the productstream may be less that 1.0%.

In other embodiments of the present invention, a weight percentage ofnormal paraffins in the product stream from the isomerization zone isapproximately 24%. Additionally, a weight percentage of methane in theproduct stream is less that 0.5%.

A weight ratio of normal paraffins to methane in a product stream fromthe isomerization zone in one or more embodiments may be at least 75, orat least 100, or approximately 114.

Additional embodiments and details of the present invention are setforth in the following detailed description of the invention.

DETAILED DESCRIPTION OF THE DRAWINGS

The drawings are simplified process diagrams in which:

FIG. 1 shows a process flow diagram of a process according to oneembodiment of the present invention;

FIG. 2 shows a process flow diagram of a process according to anotherembodiment of the present invention; and,

FIG. 3 shows a process flow diagram of a process according to anotherembodiment of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

As mentioned above, it has been discovered that the conversion to normalparaffins in the isomerization zone can be increased by removing aportion of the C₆ cyclic hydrocarbons, such as cyclohexane,methyl-cyclopentane, and benzene, in the stream passing into theisomerization zone. Specifically, when the amount of C₆ cyclichydrocarbons in the stream has been reduced, disproportionationreactions occur which lead to increased amounts of valuable C₃hydrocarbons and C₄ hydrocarbons, as well as increases in the per passconversion of the iso-paraffin hydrocarbons in the feed. The productsfrom the disproportionation reactions undergo isomerization reactionsleading to an increase in yields of normal paraffins according to thevarious embodiments of the present invention.

The ability to increase the conversion rate per pass of iso-paraffinhydrocarbons to normal paraffins in the absence of C₆+ cyclichydrocarbons in the isomerization zone is further enhanced by alsocontrolling both the temperature and the amount of hydrogen in theisomerization zone. Controlling of these variables will allow morefavorable ratios of the desired normal paraffin products over theundesired production of low value methane. The production of desiredproducts over the production of methane leads to increased value in thedownstream processing of these products such as cracking to lightolefins in a steam cracker. Accordingly, for processes of the presentinvention, the increased conversion of the iso-paraffin hydrocarbons inthe isomerization zone leads to smaller recycle streams, use of smallerequipment, lower capital expenditures, and lower operating expenses.

A first embodiment of the present invention is shown in FIG. 1, in whicha feed stream 10 is passed into a first separation zone 12. The feedstream 10 is preferably hydrotreated naphtha comprising C₅+ hydrocarbons(meaning hydrocarbons having five or more carbon atoms).

The first separation zone 12 may include a separator column 14, such asa fractionation column. As will be appreciated, the depiction of column14 is simplified as all the auxiliary operational components, such ascontrols, trays, condenser and reboiler, may be of conventional design.The feed stream 10, or multiple feed streams, can be fed into the column14 at different locations if appropriate. The column 14 will typicallycontain conventional vapor-liquid contacting equipment such as trays orpacking. The type of tray and design details such as tray type, trayspacing and layout may vary within the column 14.

The column 14 will separate the feed stream 10 into an overhead stream16 and a bottoms stream 18. The overhead stream 16 may comprise C₅hydrocarbons and iC₆ hydrocarbons. Since at least a portion of the C₆cyclic hydrocarbons have been removed from the portion of the feedstream 10 in the overhead stream 16, the overhead stream 16 will be a C₆cyclic hydrocarbons lean stream. The bottoms stream 18 may comprisen-hexane, C₆ cyclic hydrocarbons, and C₇+ hydrocarbons. Furthermore,depending on the operating conditions of the column 14, the bottomsstream 18 may also contain some small amounts of iC₆ hydrocarbons, suchas 3-methylpentane.

The bottoms stream 18 may be passed to various other zones, such as, forexample: to saturation and then to a steam cracker; to a reformer andthen to an aromatic complex; to saturation, then to a ring operatingreactor, and then to a steam cracker; or a combination of the foregoing.The further processing of the bottoms stream 18 is not necessary for theunderstanding and practicing of the present invention.

Returning to FIG. 1, the overhead stream 16 from the first separationzone 12 may be passed to a second separation zone 20. The secondseparation zone 20 provides at least one stream 22 that is rich in iC₅hydrocarbons, iC₆ hydrocarbons, or both. In a preferred embodiment, thesecond separation zone 20 comprises at least two columns 24, 26. Thesetwo columns 24, 26 may also be fractionation columns.

The first column 24 in the second separation zone 20 may receive theoverhead stream 16 from the first separation zone 12. In this embodimentof the present invention, the first column 24 separates the overheadstream 16 into three streams, and thus may comprise a divided wallcolumn. The three streams produced by the first column 24 are anoverhead stream 28, an intermediate stream 30, and a bottoms stream 32.

The overhead stream 28 comprises C₅ hydrocarbons. The intermediatestream 30 comprises iC₆ hydrocarbons. The bottoms stream 32 comprises C₆cyclic hydrocarbons, C₇+ hydrocarbons and n-hexane which either were notseparated out in the first separation zone 12 or which were formed inthe isomerization zone.

The bottoms stream 32 may be passed to various other zones, such as, forexample a steam cracker. The further processing of the bottoms stream 32is not necessary for the understanding and practicing of the presentinvention.

The intermediate stream 30 has a high concentration of iC₆ hydrocarbons,compared to the concentration of iC₆ hydrocarbons in the feed stream 10.Thus, the intermediate stream 30 is considered an iC₆ hydrocarbon richstream. The intermediate stream 30 may be passed to an isomerizationzone 34, discussed in more detail below.

The overhead stream 28 from the first column 24 is passed to the secondcolumn 26 in the second separation zone 20. In the second column 26, theoverhead stream 28 from the first column 24 of the second separationzone 20 is separated into an overhead stream 36 and a bottoms stream 38.The bottoms stream 38 comprises n-pentane and may be combined withbottoms stream 32 from the first column 24 in the second separation zone20 and passed to, for example, a steam cracker. The further processingof this stream 38 is not necessary for the understanding and practicingof the present invention.

The overhead stream 36 from the second column 26 in the secondseparation zone 20 comprises iC₅ hydrocarbons. Since the concentrationof iC₅ hydrocarbons in this stream 36 is higher than the concentrationof iC₅ hydrocarbons in the feed stream 10, it is an iC₅ hydrocarbon richstream. The overhead stream 36 from the second column 26 of the secondseparation zone 20 may be passed to the isomerization zone 34, discussedbelow. The overhead stream 36 may be combined with the intermediatestream 30 from the first column 24 of the second separation zone 20.Since both the iC₅ hydrocarbons and the iC₆ hydrocarbons streams 36, 30were separated from a portion of the C₆ cyclic hydrocarbons lean stream16, the amount of C₆ cyclic hydrocarbons passed to the isomerizationzone 34 is lower.

In the isomerization zone 34, iC₅ hydrocarbons and iC₆ hydrocarbons, inthe presence of hydrogen and an isomerization catalyst, are convertedinto normal paraffins. The isomerization zone 34, as is known, typicallycontains a series of reactors and a separation column. It is preferredthat both the iC₅ hydrocarbons and the iC₆ hydrocarbons streams 36, 30are passed to the same isomerization zone 34; however it is contemplatedthat two separate isomerization zones can be used.

While it is known that cracking of some of the paraffins can occur in anisomerization zone 34 to form C₁ to C₄ hydrocarbons, it has beendiscovered that the conversion of iC₅ and iC₆ hydrocarbons increasessignificantly via disproportionation reactions due to the fact that thestream(s) 36, 30 passed into the isomerization zone 34 are lean in C₆cyclic hydrocarbons. It is believed that the disproportionationreactions occur by the combination of two iso-paraffin hydrocarbonsfollowed by scission into one lighter hydrocarbon and one heavierhydrocarbon. For example, two iC₅ hydrocarbons can combine and form aniC₄ hydrocarbon and an iC₆ hydrocarbon in the presence of hydrogen. TheiC₄ hydrocarbons can further react via disproportionation to form a C₃hydrocarbon and an iC₅ hydrocarbon. A significant portion of theproduced iC₄ hydrocarbons also converts to nC₄ hydrocarbons viaisomerization reactions in the isomerization zone. A key discovery inthe present invention is the significant production of C₃ and C₄ normalparaffins via disproportionation and isomerization reactions with lowproduction of low-value undesired methane as a cracked product.

This surprising result is enabled by the use of an isomerizationcatalyst, and thus, the isomerization zone 34 can include, chloridedalumina, sulfated zirconia, tungstated zirconia or zeolite-containingisomerization catalysts. The isomerization catalyst may be amorphous,e.g. based upon amorphous alumina, or zeolitic. A zeolitic catalystwould still normally contain an amorphous binder. The catalyst maycomprise a sulfated zirconia and platinum as described in U.S. Pat. No.5,036,035 and European patent application 0 666 109 A1 or a platinumgroup metal on chlorided alumina as described in U.S. Pat. No. 5,705,730and U.S. Pat. No. 6,214,764. Another suitable catalyst is described inU.S. Pat. No. 5,922,639. U.S. Pat. No. 6,818,589 discloses a catalystcomprising a tungstated support of an oxide or hydroxide of a Group IVB(IUPAC 4) metal, preferably zirconium oxide or hydroxide, at least afirst component which is a lanthanide element and/or yttrium component,and at least a second component being a platinum-group metal component.These documents are incorporated herein for their teaching as tocatalyst compositions, isomerization operating conditions andtechniques.

Contacting within the isomerization zone 34 may be effected using thecatalyst in a fixed-bed system, a moving-bed system, a fluidized-bedsystem, or in a batch-type operation. The reactants may be contactedwith the bed of catalyst particles in upward, downward, or radial-flowfashion. The reactants may be in the liquid phase, a mixed liquid-vaporphase, or a vapor phase when contacted with the catalyst particles, witha mixed phase or vapor phase being preferred. The isomerization zone 34may be in a single reactor or in two or more separate reactors withsuitable means therebetween to insure that the desired isomerizationtemperature is maintained at the entrance to each zone. Two or morereactors in sequence enable improved isomerization through control ofindividual reactor temperatures and for partial catalyst replacementwithout a process shutdown.

Especially where a chlorided catalyst is used for isomerization, theisomerization reaction effluent can be contacted with a sorbent toremove any chloride components such as disclosed in U.S. Pat. No.5,705,730.

Returning to FIG. 1, a first stream 40 recovered from the isomerizationzone comprises C₄− hydrocarbons. This stream 40 may be sent to gastreatment, then to a steam cracker, or it may be sent to gas treatment,separation, and an iC₄ hydrocarbons stream may be sent to anotherisomerization zone. The further processing of this stream 40 is notnecessary for the understanding and practicing of the present invention.

A second stream 42 recovered from the isomerization zone 34 willcomprise C₅+ hydrocarbons, including normal paraffins. The second stream42 may be sent back through the first separation zone 12, the secondseparation zone 20, or both to separate out the normal paraffins fromthe iso-paraffins. For example, as shown in FIG. 1, the second stream 42is passed back to the first column 24 of the second separation zone 20.The normal hydrocarbons in the second stream 42 will be separated outwith the components of the C₆ cyclic hydrocarbons lean stream 16 passingthrough the second separation zone 20 and can be further processed asmentioned above.

Turning to FIG. 2, another embodiment of the present invention is shown.In this embodiment, a feed stream 100 is also passed into a firstseparation zone 102. The feed stream 100 is preferably hydrotreatednaphtha comprising C₅+ hydrocarbons.

The first separation zone 102 may include a separator column 104, suchas a fractionation column which preferably functions identically to thecolumn 14 in the embodiment shown in FIG. 1. Thus, the feed stream 100will separate into an overhead stream 106 and a bottom stream 108.However, it is also contemplated that the first separation zone 102comprises an adsorption zone (discussed below).

The overhead stream 106 may comprise C₅ hydrocarbons and iC₆hydrocarbons identical to stream 16 in the embodiment in FIG. 1. Sincethe C₆ cyclic hydrocarbons have been removed from the portion of thefeed stream 100 in the overhead stream 106, the overhead stream 106 willbe a C₆ cyclic hydrocarbons lean stream. The bottom stream 108 maycomprise n-hexane, C₆ cyclic hydrocarbons, and C₇+ hydrocarbons and asmall amount of iC₆ hydrocarbons. The bottoms stream 108 may be passedto various other zones, such as, for example: to saturation and then toa steam cracker; to a reformer and then to an aromatic complex; tosaturation, then to a ring operating reactor, and then to a steamcracker; or a combination of the foregoing. The further processing ofbottoms stream 108 is not necessary for the understanding and practicingof the present invention.

Returning to FIG. 2, the overhead stream 106 from the first separationzone 102 may be passed to a second separation zone 110. In thisembodiment of the present invention, it is contemplated that the secondseparation zone 110 comprises an adsorption zone 112.

The adsorption zone 112 can include, as is known, a single large bed ofadsorbent or in several parallel beds on a swing bed basis. However, ithas been found that simulated moving bed adsorptive separation providesseveral advantages such as high purity and recovery. Therefore, manycommercial scale petrochemical separations especially for the recoveryof mixed paraffins are performed using simulated countercurrent movingbed (SMB) technology. Further details on equipment and techniques foroperating an SMB process may be found in U.S. Pat. Nos. 3,208,833;3,214,247; 3,392,113; 3,455,815; 3,523,762; 3,617,504; 4,006,197;4,133,842; and 4,434,051, all of which are incorporated by reference intheir entirety. A different type of simulated moving bed operation whichcan be performed using similar equipment, adsorbent and conditions butwhich simulates co-current flow of the adsorbent and liquid in theadsorption chambers is described in U.S. Pat. Nos. 4,402,832 and4,498,991, which are incorporated by reference in their entirety.

Operating conditions for the adsorption chamber used in the subjectinvention include, in general, a temperature range of from about 20° C.to about 250° C. Adsorption conditions also preferably include apressure sufficient to maintain the process fluids in liquid phase;which may be from about atmospheric to about 600 psig. Desorptionconditions generally include the same temperatures and pressure as usedfor adsorption conditions. It is generally preferred that an SMB processis operated with an A:F flow rate through the adsorption zone in thebroad range of about 1:1 to 5:0.5 where A is the volume rate of“circulation” of selective pore volume and F is the feed flow rate. Thepractice of the subject invention requires no significant variation inoperating conditions or desorbent composition within the adsorbentchambers. That is, the adsorbent preferably remains at the sametemperature throughout the process during both adsorption anddesorption.

The adsorbent used in the adsorption zone 112 preferably comprisesaluminosilicate molecular sieves having relatively uniform porediameters of about 5 angstroms. This is provided by commerciallyavailable type 5A molecular sieves produced by UOP LLC.

A second adsorbent which could be used in the adsorption zone 112comprises silicalite. Silicalite is well described in the literature. Itis disclosed and claimed in U.S. Pat. No. 4,061,724 issued to Grose etal., which is incorporated by reference in its entirety. A more detaileddescription is found in the article, “Silicalite, A New HydrophobicCrystalline Silica Molecular Sieve,” Nature, Vol. 271, Feb. 9, 1978which is incorporated herein by reference for its description andcharacterization of silicalite. Silicalite is a hydrophobic crystallinesilica molecular sieve having intersecting bent-orthogonal channelsformed with two cross-sectional geometries, 6 Å circular and 5.1-5.7 Åelliptical on the major axis. This gives silicalite great selectivity asa size selective molecular sieve. Due to its aluminum free structurecomposed of silicon dioxide, silicalite does not show ion-exchangebehavior. Silicalite is also described in U.S. Pat. Nos. 5,262,144;5,276,246 and 5,292,900, which are incorporated by reference in theirentirety. These basically relate to treatments which reduce thecatalytic activity of silicalite to allow its use as an adsorbent.

The active component of the adsorbent is normally used in the form ofparticle agglomerates having high physical strength and attritionresistance. The agglomerates contain the active adsorptive materialdispersed in an amorphous, inorganic matrix or binder, having channelsand cavities therein which enable fluid to access the adsorptivematerial. Methods for forming the crystalline powders into suchagglomerates include the addition of an inorganic binder, generally aclay comprising a silicon dioxide and aluminum oxide, to a high purityadsorbent powder in a wet mixture. The binder aids in forming oragglomerating the crystalline particles. The blended clay-adsorbentmixture may be extruded into cylindrical pellets or formed into beadswhich are subsequently calcined in order to convert the clay to anamorphous binder of considerable mechanical strength. The adsorbent mayalso be bound into irregular shaped particles formed by spray drying orcrushing of larger masses followed by size screening. The adsorbentparticles may thus be in the form of extrudates, tablets, spheres orgranules having a desired particle range, preferably from about 16 toabout 60 mesh (Standard U.S. Mesh) (1.9 mm to 250 microns). Clays of thekaolin type, water permeable organic polymers or silica are generallyused as binders.

The active molecular sieve component of the adsorbent will preferably bein the form of small crystals present in the adsorbent particles inamounts ranging from about 75 to about 98-wt. % of the particle based onvolatile-free composition. Volatile-free compositions are generallydetermined at 900° C., after the adsorbent has been calcined, in orderto drive off all volatile matter. The remainder of the adsorbent willgenerally be the inorganic matrix of the binder present in intimatemixture with the small particles of the silicalite material. This matrixmaterial may be an adjunct of the manufacturing process for thesilicalite, for example, from the intentionally incomplete purificationof the silicalite during its manufacture.

Those skilled in the art will appreciate that the performance of anadsorbent is often greatly influenced by a number of factors not relatedto its composition such as operating conditions, feed stream compositionand the water content of the adsorbent. The optimum adsorbentcomposition and operating conditions for the process are thereforedependent upon a number of interrelated variables. One such variable isthe water content of the adsorbent which is expressed herein in terms ofthe recognized Loss on Ignition (LOI) test. In the LOI test the volatilematter content of the zeolitic adsorbent is determined by the weightdifference obtained before and after drying a sample of the adsorbent at500° C. under an inert gas purge such as nitrogen for a period of timesufficient to achieve a constant weight. For the subject process it ispreferred that the water content of the adsorbent results in an LOI at900° C. of less than 7.0% and preferably within the range of from 0 to4.0 wt %.

An important characteristic of an adsorbent is the rate of exchange ofthe desorbent for the extract component of the feed mixture materialsor, in other words, the relative rate of desorption of the extractcomponent. This characteristic relates directly to the amount ofdesorbent material that must be employed in the process to recover theextract component from the adsorbent. Faster rates of exchange reducethe amount of desorbent material needed to remove the extract component,and therefore, permit a reduction in the operating cost of the process.With faster rates of exchange, less desorbent material has to be pumpedthrough the process and separated from the extract stream for reuse inthe process. Exchange rates are often temperature dependent. Ideally,desorbent materials should have a selectivity equal to about 1 orslightly less than 1 with respect to all extract components so that allof the extract components can be desorbed as a class with reasonableflow rates of desorbent material, and so that extract components canlater displace desorbent material in a subsequent adsorption step.

U.S. Pat. No. 4,992,618 issued to S. Kulprathipanja, and which isincorporated by reference in its entirety, describes the use of a“prepulse” of a desorbent component in an SMB process for recoveringnormal paraffins. The prepulse is intended to improve the recovery ofthe extract normal paraffins across the carbon number range of the feed.The prepulse enters the adsorbent chamber at a point before (downstream)the feed injection point. A related SMB processing technique is the useof “zone flush.” The zone flush forms a buffer zone between the feed andextract bed lines to keep the desorbent from entering the adsorptionzone. While the use of a zone flush requires a more complicated, andthus more costly rotary valve, the use of zone flush is preferred in theadsorption zones when high purity extract product are desired. Inpractice, a quantity of the mixed component desorbent recovered overheadfrom the extract and raffinate columns may be passed into a separatesplitter column. A high purity stream of the lower strength component ofthe mixed component desorbent is recovered and used as the zone flushstream. Further information on the use of dual component desorbents andon techniques to improve product purity such as the use of flush streamsmay be obtained from U.S. Pat. Nos. 3,201,491; 3,274,099; 3,715,409;4,006,197 and 4,036,745 which are incorporated herein by reference intheir entirety for their teaching on these aspects of SMB technology.

It has become customary in the art to group the numerous beds in the SMBadsorption chamber(s) into a number of zones. Usually the process isdescribed in terms of 4 or 5 zones. First contact between the feedstream and the adsorbent is made in Zone I, the adsorption zone. Theadsorbent or stationary phase in Zone I becomes surrounded by liquidwhich contains the undesired isomer(s), that is, the raffinate. Thisliquid is removed from the adsorbent in Zone II, referred to as apurification zone. In the purification zone the undesired raffinatecomponents are flushed from the void volume of the adsorbent bed by amaterial which is easily separated from the desired component byfractional distillation. In Zone III of the adsorbent chamber(s) thedesired isomer is released from the adsorbent by exposing and flushingthe adsorbent with the desorbent (mobile phase). The released desiredisomer and accompanying desorbent are removed from the adsorbent in theform of the extract stream. Zone IV is a portion of the adsorbentlocated between Zones I and III which is used to segregate Zones I andIII. In Zone IV desorbent is partially removed from the adsorbent by aflowing mixture of desorbent and undesired components of the feedstream. The liquid flow through Zone IV prevents contamination of ZoneIII by Zone I liquid by flow cocurrent to the simulated motion of theadsorbent from Zone III toward Zone I. A more thorough explanation ofsimulated moving bed processes is given in the Adsorptive Separationsection of the Kirk-Othmer Encyclopedia of Chemical Technology at page563. The terms “upstream” and “downstream” are used herein in theirnormal sense and are interpreted based upon the overall direction inwhich liquid is flowing in the adsorbent chamber. That is, if liquid isgenerally flowing downward through a vertical adsorbent chamber, thenupstream is equivalent to an upward or higher location in the chamber.

In an SMB process the several steps e.g. adsorption and desorption, arebeing performed simultaneously in different parts of the mass ofadsorbent retained in the adsorbent chamber(s) of the process. If theprocess was being performed with two or more adsorbent beds in a swingbed system then the steps may be performed in a somewhat interruptedbasis, but adsorption and desorption will most likely occur at the sametime.

Returning to FIG. 2, a first stream 114 and a second stream 116 arerecovered from the adsorption zone 112. The first stream 114 comprisesnormal paraffins and may be sent to, for example, a stream cracker.

The second stream 116 recovered from the adsorption zone 112 comprisesiso-paraffins, or is rich in iC₅ and iC₆ hydrocarbons. This stream 116is passed to an isomerization zone 118. As with the isomerization zone34 in the embodiment shown in FIG. 1, in the isomerization zone 118 ofthis embodiment, iC₅ and iC₆ hydrocarbons, in the presence of hydrogenand an isomerization catalyst, are converted into normal paraffins.

As the stream 116 introduced into the isomerization zone 118 is lean inC₆ cyclic hydrocarbons, there is a surprising and unexpected increase inthe yields to desirable normal paraffins believed to be produced viadisproportionation and isomerization reactions. The specifics of thisisomerization zone 118 are the same as discussed above, and thus, areincorporated herein to the discussion of this embodiment.

At least two streams 120, 122 may also be recovered from theisomerization zone 118. The first stream 120 comprises C₄− hydrocarbons.This stream 120 may be sent to gas treatment, then to a steam cracker,or it may be sent to gas treatment, separation, and an iC₄ hydrocarbonsstream may be sent to another isomerization zone. The further processingof this stream 120 is not necessary for the understanding and practicingof the present invention.

The second stream 122 recovered from the isomerization zone 118 willagain comprise C₅+ hydrocarbons, including normal paraffins. The secondstream 122 may be sent back though the first separation zone 12, thesecond separation zone 110, or both to separate out the normal paraffinsfrom the iso-paraffins. In one embodiment, the second stream 122 ispassed back to the first separation zone 102. It may or may not becombined with fresh feed stream 100 entering the first separation zone102. Thus, the normal hexane will be separated out in the firstseparation zone 102, while the normal pentane will be separated out inthe second separation zone 110. The normal paraffins will be passedalong to further processing units, as discussed above.

In a third embodiment according to the present invention as shown inFIG. 3, a feed stream 200 is passed into a first separation zone 202.The feed stream 200 is preferably hydrotreated naphtha comprising C₅+hydrocarbons.

The first separation zone 202 may also include a separator column 204,such as a fractionation column. This column 204 will separate the feedstream 200 into an overhead stream 206, an intermediate stream 208, anda bottom stream 210.

The overhead stream 206 may comprise C₅ hydrocarbons and iC₆hydrocarbons. The intermediate stream 208 may comprise n-hexane and C₆cyclic hydrocarbons. The bottoms stream 210 may comprise C₇+hydrocarbons. The bottoms stream 210 may be passed to various otherzones, such as, for example: to saturation and then to a steam cracker;to a reformer and then to an aromatic complex; to saturation, then to aring operating reactor, and then to a steam cracker; or a combination ofthe foregoing. The further processing of bottoms stream 210 is notnecessary for the understanding and practicing of the present invention.

The overhead stream 206 may be passed to a second separation zone 212.Since the C₆ cyclic hydrocarbons have been removed from the portion ofthe feed stream 200 in the overhead stream 206, the overhead stream 206will be a C₆ cyclic hydrocarbons lean stream.

It is contemplated that the second separation 212 comprise one or morecolumns, such as the second separation zone 20 in the embodiment shownin FIG. 1, or it may comprise an adsorption zone, such the secondseparation zone 110 shown in FIG. 2. Accordingly, the portions of thoseembodiments are incorporated herein.

A first stream 214 from the second separation zone 212 is rich iniso-paraffins and is passed to an isomerization zone 216. A secondstream 218 from the second separation zone 212 is rich in normalparaffins may be passed to further processing zones. The isomerizationzone 216 and the processing of the second stream 218 from the secondseparation zone may be the same as discussed above with respect to theother embodiments of the present invention.

Returning to the first separation zone 202 in FIG. 3, the differencebetween this embodiment and the previously discussed embodiments is theintermediate stream 208 which is passed to a ring opening reaction zone220. In the ring opening reactor zone 220, the cyclic hydrocarbons, inthe presence of a catalyst, are converted into straight chainhydrocarbons. Typically, such reactions occur in a ring opening reactor222. Such ring opening reactors are known, for example, as disclosed inU.S. Pat. Pub. No. 2005/0101814, incorporated herein by reference.

The products of the ring opening reactor 222, which can include methaneto C₇+ hydrocarbons, may be separated into a C₄− hydrocarbon stream 224,a C₅ hydrocarbon and C₆ hydrocarbon stream 226, and a C₆ cyclichydrocarbons and C₇+ hydrocarbons stream 228. The C₆ cyclic hydrocarbonsand C₇+ hydrocarbons stream 228 may be combined with the bottoms stream210 from the first separation zone 202. The C₄− hydrocarbon stream 224may be passed to further processing units or zones. The C₅ hydrocarbonand C₆ hydrocarbon stream 226 may be combined with the overhead stream206 of the first separation zone 202, and thus passed to the separationzone 212 and isomerization zone 216.

The above described embodiments are merely exemplary of the presentinvention. Other processes for producing a C₆ cyclic hydrocarbons leanstream may be utilized.

In accordance with the present invention, the isomerization conditionsin the isomerization zone include average reactor temperatures usuallyranging from about 40° C. to 250° C. The average reactor temperature isdefined as the integrated average temperature over the catalyst bed in areactor. The reactor operating pressures generally range from about 100kPa to 10 MPa absolute. Liquid hourly space velocities (LHSV) range fromabout 0.2 to about 25 volumes of isomerizable hydrocarbon feed per hourper volume of catalyst.

Hydrogen is admixed with or remains with the isomerization feed to theisomerization zone to provide a mole ratio of hydrogen to hydrocarbonfeed of from about 0.01 to 20. The hydrogen may be supplied totally fromoutside the process or supplemented by hydrogen recycled to the feedafter separation from isomerization reactor effluent. Light hydrocarbonsand small amounts of inerts such as nitrogen and argon may be present inthe hydrogen. Water should be removed from hydrogen supplied fromoutside the process, preferably by an adsorption system as is known inthe art.

Over a range of LHSV (1 to 5 hr⁻¹) for a feed stream which is lean inC₆+ cyclic hydrocarbons in the presence of a chlorided-aluminaisomerization catalyst, more favorable ratios of the yield of normalparaffins, which include ethane, propane, nC₄, nC₅ and nC₆ hydrocarbons,to the yield of methane are obtained by regulating the outlet hydrogento hydrocarbon feed mole ratio (H₂/HC) to less than 0.7, preferably, toless than about 0.2, while operating at average reactor temperaturesgreater then about 176.6° C. (350° F.), preferably greater than about190.5° C. (375° F.). Similarly for other isomerization catalyst types,the temperatures and H₂/HC ratios can be set to obtain favorable ratiosof desired normal paraffins to undesired methane.

Furthermore, it is contemplated that an amount of C₆ cyclic hydrocarbonspassed to the isomerization zone is adjusted. This is believed to allowfor control of at least the disproportionation reactions and thecustomization of the product streams. For example, an operatingparameter of the various separation zones may be controlled so that anamount of C₆ cyclic hydrocarbons does reach the isomerization zone.Additionally, and alternatively, a C₆ cyclic hydrocarbons rich streammay be introduced into the steams passing into the isomerization zone.

To demonstrate the advantages of the present invention, achlorided-alumina catalyst that contained platinum was loaded andoperated under isomerization conditions of 3.1 MPa (450 psig), with a0.06 outlet hydrogen to hydrocarbon feed (H₂/HC) mole ratio, and at arate of 2 h⁻¹ LHSV with an average temperature of approximately 174.4°C. (346° F.).

Feed A, which comprised 97% iC₅ and 3% nC₅ hydrocarbons and contained noC₆ cyclic hydrocarbons, was processed. As shown in Table 1 for ProductA, significant quantities of C₆ and C₄ hydrocarbons were made viadisproportionation reactions, such as 2iC₅→iC₄+iC₆. The normal paraffinsthat are produced are a result of isomerization reactions such as iC₄4⇄nC₄ which are limited by equilibrium. The C₃ hydrocarbons that areproduced are a result of the disproportionation reaction 2iC₄→C₃+iC₅.

TABLE 1 COMPONENTS FEED PRODUCT (wt %) A A H₂ — 0.2 C₁ — 0.1 C₂ — 0.1 C₃— 1.3 iC₄ 0.0 9.6 nC₄ 0.0 3.8 iC₅ 96.8 53.4 nC₅ 3.1 18.0 iC₆ 0.0 11.8nC₆ 0.0 1.7 Cyclopentane (CP) 0.0 0.0 Methylcyclopentane (MCP) 0.0 0.0Cyclohexane (CH) 0.0 0.0 Benzene (BZ) 0.0 0.0 C₇+ 0.0 0.2 SUM 100.0100.0 iC₅ converted (%) — 44.9 nC₅ + nC₆ 3.1 19.8 C₂, C₃, nC₄, 3.1 24.9nC₅, nC₆

As can be seen in Table 1, the sum of C₂ to C₆ normal paraffins washigher at 24.9 wt % when including the products from disproportionationas compared to 19.8 wt % yield when only including the nC₅+nC₆hydrocarbons from isomerization.

In a second experiment, another chlorided-alumina catalyst thatcontained platinum was loaded and operated under isomerizationconditions of 3.1 MPa (450 psig), a 0.2 outlet H₂/HC mole ratio and arate 2 h⁻¹ LHSV with an average temperature at about 176.6° C. (350°F.). Feed B was rich in iC₅ and iC₆ hydrocarbons and contained 1.46%cyclopentane (CP). Feed C was similar to Feed B with 1.42 wt %cyclopentane but also contained 1.29 wt % cyclohexane (a C₆ cyclichydrocarbon). With the C₆ cyclic hydrocarbon in Feed C, as shown inTable 2, the amount of C₃ and C₄ hydrocarbons were greatly reduced(compare Product C vs. Product B). This demonstrates that the presenceof the C₆ cyclic hydrocarbon significantly decreased thedisproportionation reactions.

TABLE 2 FEED PRODUCT FEED PRODUCT B B C C COMPONENTS (wt %) H₂ 0.5 0.5C₁ 0.2 0.1 C₂ 0.3 0.1 C₃ 3.6 1.0 iC₄ 0.0 9.2 0.0 3.2 nC₄ 0.0 4.5 0.0 1.0iC₅ 56.1 34.1 56.1 39.8 nC₅ 1.8 11.5 1.8 13.3 iC₆ 39.4 31.4 38.2 34.6nC₆ 1.1 4.6 1.1 5.1 CP, MCP, CH, BZ 1.5 0.0 2.8 0.8 C₇+ 0.0 0.2 0.0 0.7SUM 100.0 100.0 100.0 100.0 iC₅ + iC₆ — 31.4 — 21.2 converted (%) C₂,C₃, nC₄, 2.9 24.4 2.9 20.4 nC₅, nC₆ CYCLICS CP 1.46 0.00 1.42 0.56 MCP0.06 0.01 0.06 0.12 CH 0.03 0.01 1.29 0.08 BZ 0.00 0.00 0.00 0.00 C₇+Cyclics 0.00 0.03 0.00 0.37 TOTAL CYCLICS 1.54 0.05 2.77 1.12 Cyclics —97.0 — 59.5 Converted (%) CP Conversion (%) — 100.0 — 60.9

In addition, as shown in Table 2, the iC₅ and iC₆ hydrocarbon conversionand the ring opening conversion were lower when C₆ cyclic hydrocarbonswere present in the feed. It can additionally be observed that withoutthe C₆ cyclic hydrocarbons, the sum of C₂ to C₆ normal paraffins inProduct B was 24.4 wt %. However, with the C₆ cyclic hydrocarbon in thefeed, the sum of C₂ to C₆ normal paraffins in Product C was less, 20.4wt %.

Thus, as shown, the iC₅ and iC₆ hydrocarbon conversions and the normalparaffin yields can be increased in the isomerization zone by removingor significantly reducing the C₆ cyclic hydrocarbons from the streampassed into the isomerization zone.

In a third experiment, a chlorided-alumina catalyst that containedplatinum was loaded and operated under isomerization conditions of 3.1MPa (450 psig), at a rate of 2 h⁻¹ LHSV with an average catalyst bedtemperature of approximately 179.4° C. (355° F.). As shown in Table 3, afeed stream comprising C₅ and C₆ hydrocarbons and 1.5 wt % cyclopentane,but no C₆+ cyclic components was used (Feed D).

TABLE 3 Feed Product Product D D1 D2 Conditions Outlet H₂/HC 0.2 0.7COMPONENTS (wt %) H₂ 0.5 1.8 C₁ 0.2 1.5 C₂ 0.3 2.0 C₃ 3.6 10.0 iC₄ 0.09.2 17.8 nC₄ 0.0 4.5 12.1 iC₅ 56.1 34.1 25.8 nC₅ 1.8 11.5 8.8 iC₆ 39.431.4 17.7 nC₆ 1.1 4.6 2.6 CP 1.5 0.0 0.0 C₇+ 0.0 0.2 0.0 SUM 100.0 100.0100.0 C₂, C₃, nC₄, 2.9 24.4 35.4 nC₅, nC₆ (nC₂-C₆)/C₁, — 122 24 wt ratio

Table 3 shows that with a 0.2 outlet H₂/HC feed mole ratio, the yield ofnormal paraffins was 24.4 wt % (sum of desired components C₂, C₃, nC₄,nC₅, nC₆) while the yield of methane (which is not a desired product)was 0.2 wt % (Product D1). However, when an outlet H₂/HC mole ratio of0.7 was used, the yield of normal paraffins increased to 35.4 wt % butthe methane also increased substantially to 1.5 wt % (Product D2). Theweight ratio of desired products divided by undesired product decreasedfrom 122 to 24 by raising the outlet H₂/HC mole ratio from 0.2 to 0.7.

Therefore, to avoid excessive methane production, it is mostadvantageous to operate at lower outlet H₂/HC mole ratios, preferablylower that 0.7, and most preferably less than about 0.2. Using suchisomerization conditions leads to higher selectivity to normal paraffinyields over methane yields.

In a fourth experiment, a chlorided-alumina catalyst that containedplatinum was loaded and operated under isomerization conditions of 3.1MPa (450 psig), at a rate of 2 h⁻¹ LHSV with average catalyst bedtemperatures ranging from 175° C. (347° F.) to 203.9° C. (399° F.). Asshown in Table 4, a feed stream comprising C₅ and C₆ hydrocarbons withno C₆+ cyclic components was used (Feed E).

TABLE 4 FEED PRODUCT PRODUCT PRODUCT E E1 E2 E3 CONDITIONS Temperature °C. 175 190.5 203.9 Outlet H₂/HC 0.2 0.1 0.03 COMPONENTS (wt %) H₂ 0.50.3 0.1 C₁ 0.2 0.1 0.1 C₂ 0.2 0.2 0.1 C₃ 4.2 4.8 5.2 iC₄ 0.0 13.1 14.715.6 nC₄ 0.0 5.6 7.1 7.9 iC₅ 55.5 30.9 29.5 28.2 nC₅ 1.9 10.5 10.7 10.8iC₆ 41.9 29.1 26.9 25.6 nC₆ 0.5 4.2 4.2 4.2 CP 0.0 0.0 0.1 0.1 C₇+ 0.21.4 1.6 2.1 SUM 100.0 100.0 100.0 100.0 C2, C₃, nC₄, 2.3 24.8 26.9 28.2nC₅, nC₆ (nC₂-C₆)/C₁, — 124 269 282 wt ratio

Table 4 shows that raising the average catalyst bed temperature from175° C. (347° F.) to 203.9° C. (399° F.) while maintaining low outletH₂/HC mole ratios (below about 0.2), leads to an increase in the desirednormal paraffin products (C₂, C₃, nC₄, nC₅, nC₆) with a decrease in themethane yield. The weight ratio of desired normal paraffin productsdivided by undesired methane product was highest for the case with203.9° C. (399° F.) average catalyst bed temperature and 0.03 outletH₂/HC mole ratio (Product E3).

In a fifth experiment, a chlorided-alumina catalyst that containedplatinum was loaded and operated under isomerization conditions of 3.1MPa (450 psig), at a rate of 2 h⁻¹ LHSV with average catalyst bedtemperatures of about 205° C. (401° F.). As shown in Table 5, a feedstream comprising C₅ and C₆ hydrocarbons with 4.5 wt % cyclopentane (CP)and no C₆+ cyclic components was used (Feed F).

Table 5 shows for operations at about 205° C. (about 400° F.), reducingthe outlet H₂/HC mole ratio from 0.2 to 0.1 resulted in an increase inthe ratio of normal paraffins to methane from 65 (Product F2) to 139(Product F1). This demonstrates that the outlet H₂/HC mole ratio of lessthan about 0.2 leads to higher selectivity to normal paraffin yieldsover methane yields.

TABLE 5 FEED PRODUCT PRODUCT F F1 F2 CONDITIONS Outlet H₂/HC 0.1 0.2COMPONENTS (wt %) H₂ 0.2 0.5 C₁ 0.2 0.5 C₂ 0.2 0.5 C₃ 4.7 7.6 iC₄ 0.010.8 13.8 nC₄ 0.0 6.3 9.6 iC₅ 56.1 32.3 29.3 nC₅ 1.8 12.1 11.1 iC₆ 36.527.2 22.7 nC₆ 1.0 4.5 3.9 CP 4.5 0.9 0.2 C₇+ 0.0 0.6 0.2 SUM 100.0 100.0100.0 C₂, C₃, nC⁴, 2.8 27.8 32.6 nC₅, nC₆ (nC₂-C₆)/C₁, — 139 65 wt ratio

Finally, based upon the above, it is also contemplated that similareffects are expected with respect to C₇+ cyclic hydrocarbons as thoseobserved for the C₆ cyclic hydrocarbons.

It should be appreciated and understood by those of ordinary skill inthe art that various other components such as valves, pumps, filters,coolers, etc. were not shown in the drawings as it is believed that thespecifics of same are well within the knowledge of those of ordinaryskill in the art and a description of same is not necessary forpracticing or understating the embodiments of the present invention.

While at least one exemplary embodiment has been presented in theforegoing detailed description of the invention, it should beappreciated that a vast number of variations exist. It should also beappreciated that the exemplary embodiment or exemplary embodiments areonly examples, and are not intended to limit the scope, applicability,or configuration of the invention in any way. Rather, the foregoingdetailed description will provide those skilled in the art with aconvenient road map for implementing an exemplary embodiment of theinvention, it being understood that various changes may be made in thefunction and arrangement of elements described in an exemplaryembodiment without departing from the scope of the invention as setforth in the appended claims and their legal equivalents.

What is claimed is:
 1. A process for increasing the yield of anisomerization zone, the process comprising: separating a portion of C₆cyclic hydrocarbons from a naphtha stream comprising C₅+ hydrocarbons toprovide a C₆ cyclic hydrocarbons lean stream; separating iC₅hydrocarbons and iC₆ hydrocarbons from the C₆ cyclic hydrocarbons leanstream; passing at least one stream being rich in iC₅ hydrocarbons, iC₆hydrocarbons, or both to an isomerization zone; and, maintaining anoutlet molar ratio of hydrogen to feed hydrocarbon in the isomerizationzone to be less than about 0.2.
 2. The process of claim 1 furthercomprising: maintaining an average temperature of the isomerization zoneto be at least approximately 176° C.
 3. The process of claim 1 furthercomprising: maintaining an average temperature of the isomerization zoneto be at least approximately 190° C.
 4. The process of claim 2 wherein aweight percentage of normal paraffins in a product stream from theisomerization zone is at least 20%.
 5. The process of claim 4 wherein aweight percentage of methane in the product stream is less than 1.0%. 6.The process of claim 4 wherein a weight percentage of methane in theproduct stream is less than 0.5%.
 7. The process of claim 1 wherein aweight ratio of normal paraffins to methane in a product stream from theisomerization zone is at least
 75. 8. The process of claim 1 wherein aweight ratio of normal paraffins to methane in a product stream from theisomerization zone is at least
 100. 9. The process of claim 6 wherein aweight ratio of normal paraffins to methane in a product stream from theisomerization zone is at least
 100. 10. A process for increasing theyield of an isomerization zone, the process comprising: removing C₆cyclic hydrocarbons from at least one stream comprising iC₅hydrocarbons, iC₆ hydrocarbons, or both, to provide a C₆ cyclichydrocarbons lean stream; and, passing the C₆ cyclic hydrocarbons leanstream an isomerization zone; wherein the isomerization zone has anoperating temperature of at least 176° C. and wherein an outlet molarratio of hydrogen to hydrocarbon in the isomerization zone is less thanabout 0.7.
 11. The process of claim 10 wherein the operating temperatureis at least 190° C.
 12. The process of claim 10 wherein the outlet molarratio of hydrogen to hydrocarbon in the isomerization zone is less thanabout 0.2.
 13. The process of claim 12 wherein a weight ratio of normalparaffins to methane in a product stream from the isomerization zone isat least
 75. 14. The process of claim 10 wherein a weight ratio ofnormal paraffins to methane in a product stream from the isomerizationzone is at least
 100. 15. The process of claim 14 wherein a weightpercentage of normal paraffins in the product stream from theisomerization zone is at least 20%.
 16. The process of claim 15 whereina weight percentage of methane in the product stream is less that 1.0%.17. The process of claim 10 wherein a weight percentage of normalparaffins in the product stream from the isomerization zone isapproximately 24%.
 18. The process of claim 17 wherein a weightpercentage of methane in the product stream is less that 0.5%.
 19. Theprocess of claim 18 wherein a weight ratio of normal paraffins tomethane in a product stream from the isomerization zone is at least 100.20. The process of claim 18 wherein a weight ratio of normal paraffinsto methane in a product stream from the isomerization zone isapproximately 114.